Production of high-octane, unleaded motor fuel

ABSTRACT

A charge stock containing kerosene boiling range hydrocarbons is converted into a high-octane motor fuel which does not require the use of metal-containing additives otherwise needed for suitable anti-knock characteristics. The process involves a combination of hydrocracking and catalytic reforming, and is effected in a manner which significantly increases the quantity of normally liquid motor fuel product. The novel form of hydrocracking results in a product predominantly comprising naphthenic hydrocarbons and highly-branched paraffins, the latter being rich in isobutane. Following separation to recover the isobutane, catalytic reforming is utilized to dehydrogenate the naphthenic compounds, and dehydrocyclisize paraffinic hydrocarbons, to produce an aromatic concentrate.

Donaldson et al.

[111 g Jan. 22, 1970 Primary Examiner-Herbert Levine Attorney, Agent, or Firm-.lames 1R. Hoatson et al.

A charge stock containing kerosene boiling range hydrocarbons is converted into a high-octane motor fuel which does not require the use of metal-containing additives otherwise needed for suitable anti-knock characteristics. The process involves a combination of hydrocracking and catalytic reforming, and is effected in a manner which significantly increases the quantity of normally liquid motor file] product. The novel form of hydrocracking results in a product predominantly comprising naphthenic hydrocarbons and highlybranched paraffins, the latter being rich in isobutane. Following separation to recover the isobutane, catalytic reforming is utilized to dehydrogenate the naphthenic compounds, and dehydrocyclisize paraffinic hydrocarbons, to produce an aromatic concentrate.

10 Claims, 1 Drawing Figure PRODUCTKUN 0F lllClll-OCTANE,

UNLEADED MOTOR FUEL [75] Inventors: George R. Donal, Barrington;

Ernest 1L. Polllitzor, Slrokie, both of Ill.

[73] Assignee: Universal Gil Products Company, Des Plaines, Ill.

[22] Filed: Nov. 21, 1972 [21] Appl. No.: 308,353

[52] 10.8. Cl 208/60, 208/80, 208/93 [51] int. Cl Clllg 39/00 [58] Field of Search 208/60, 79, 80, 93

[56] Relerences Cited UNITED STATES PATENTS 2,703,308 3/1955 Oblad et al. 208/60 3,l59,565 12/1964 Kimberlin et al... 208/80 3,l72,842 3/1965 Paterson 208/80 3,535,226 10/1970 .laffe 208/60 [5 Fracl/ana/ar I it i Kerosene Charge F ractian afar Ca fa/yf/c Salve/1f /-Rafarming ffxfraclion PRODUCTION OF I'IIGll'I-OCTANE, UNLEADIED MOTOR FUEL APPLICABILITY OF INVENTION The present invention involves a multiple-stage combination process for the conversion of kerosene boiling range hydrocarbons, to produce a high-octane motor fuel. The process is effected in a manner which produces an aromatic concentrate and isomeric paraffins, the latter being predominantly isobutane. More specifically, the inventive concept herein described encompasses an integrated refinery process for producing a high-octane, unleaded motor fuel gasoline pool. Feed stocks to the present process can include both naphtha and kerosene boiling range hydrocarbons i.e., a pentane to 600 F. charge stock.

Aromatic hydrocarbons, principally benzene, toluene, ethylbenzene and the xylene isomers, are required in large quantities to satisfy an ever-increasing demand for a wide variety of petrochemicals. For example, large quantities of benzene are hydrogenated to produce cyclohexane for use in the manufacture of nylon; toluene is'often used as a solvent and as the starting material for various medicines, dyes, perfumes, etc. Possibly, the principal utilization of aromatic hydrocarbons resides in gasoline blending in view of their exceedingly high research octane blending values. For example, benzene has a clear research octane blending value of about 99, while toluene and the other alkyl-substituted aromatics have values exceeding 100.

Isobutane finds wide-spread use in organic synthesis, as a refrigerant and as an aerosol propellant, etc. Other uses include conversion to isobutenes for the subsequent production of butyl rubber, copolymer resins with butadiene, acrylonitrile, etc. In accordance with one embodiment of the present invention, the multiplestage process is integrated into an overall refinery scheme for the production of a high-octane, unleaded gasoline pool. The aromatic concentrate is directly employed in the gasoline pool, while the isobutane is subjected to alkylation, with a suitable olefinic hydrocarbon, the normally liquid alkylate product being recovered as a part of the unleaded gasoline pool.

Relatively recent investigations into the causes and cures of environmental pollution have shown that more than half of the violence perpetrated upon the .atmosphere stems from vehicular exhaust, consisting primarily of unburned hydrocarbons and carbon monoxide. These investigations have brought about the development of catalytic converters which, when installed within the automotive exhaust system, are capable of converting more than 90.0 percent of the noxious components into innocuous materials. While developing these catalytic converters, it was learned that the efficiency of conversion, as well as the stability of the selected catalytic composite were severely impaired when the exhaust fumes resulted from the combustion of lead-containing motor fuel. When compared to the operation of the converter during the combustion of clear, unleaded gasoline, both the conversion of noxious components and the stability of the catalytic composites decreased as much as 50.0 percent. Therefore, it is being recognized throughout the petroleum industry, as well as in major gasoline-consuming countries, that suitable motor fuel must ultimately be produced for consumption in current internal combustion engines without requiring the addition of lead to increase the octane rating and thereby enhance the anti-knock properties. Also being recognized is the fact that unburned hydrocarbons and carbon monoxide are not the only extremely dangerous pollutants being discharged by way of vehicular exhaust. Japan has recently experienced an increase in the incidence of lead poisoning, and has enacted legislation which reduces significantly the quantity of lead permitted in motor fuel intended for consumption in that country.

A natural-flowing consequence of the removal of lead from motor fuel gasoline, in addition to others, resides in the fact the petroleum refining operations and techniques will necessarily undergo modification in order to produce voluminous quantities of a highoctane, unleaded motor fuel in an economically attractive fashion. As a direct result, the demand for naphtha boiling range charge stocks w i.e., C to 400 F. will increase to the extent that heavier feeds will be converted to naphtha in ever-increasing quantities. One well-known and well-documented refining process, capable of significantly improviing the octane rating of naphtha boiling range fractions, is the catalytic reforming process. In such a process, the principal ocatne-improving reactions are naphthene dehydrogenation, naphthene dehydroisomerization, paraffin isomerization, dehydrocyclization and hydrocracking. Naphthene dehydrogenation is an extremely rapid reaction, and constitutes the principal octane-improving reaction. With respect to a five-member ring alkylnaphthene, it is necessary first to effect isomerization to produce a six-membered ring naphthene, followed by dehydrogenation to an aromatic hydrocarbon. Paraffin dehydrocyclization is achieved through the conversion of straight-chain paraffins having at least 6 carbon atoms per molecule. The degree of paraffin aromatization is limited by virtue of the fact that the aromatic concentration increases as the reactants traverse the reforming reaction zone. Unreacted, relatively lowoctane paraffins are, therefore, present in the reformed product effluent and effectively reduce the overall octane rating thereof. In the past, these components could be tolerated due to a high susceptibility to the addition of lead.

One disadvantage of the catalytic reforming process stems from the stringent limitation placed upon the end boiling point of the reformer feed stock. Universally, charge stocks intended for conversion into high octane motor fuel have been limited to the naphtha boiling range; that is, to those hydrocarbon mixtures having an end boiling point below about 400 F. or 410 F. Such a restriction is necessary in order to maintain acceptable catalyst activity and stability. Our invention affords the opportunity to process charge stocks containing kerosene boiling range material up to a temperature of about 550 F. to 600 F.

The present invention is in part founded upon recognition of the fact that the higher paraffins i.e., C,,-, (1 and C -paraffins, etc. are more easily converted to aromatic hydrocarbons, via dehydrocyclization at relatively low operating severity levels, whereas the lower paraffins require a relatively higher severity level; the latter is accompanied by the adverse effect of attendant rampant cracking. The novel hydrocracking system, forming an integral part of the present combination process, retains virtually 100.0 percent of the cyclic hydrocarbons in the feed stock; however, al-

kyl-substituted aromatic compounds can lose their alkyl groups. These two considerations contribute to an increase in the production of light gaseous hydrocarbons. However, the hydrocracking system utilized herein effects the conversion of polynuclear aromatics and naphthenes into alkyl-substituted mononuclear naphthenes which can be dehydrogentated to aromatics in the reforming reaction zone. Additionally, the heavier paraffinic hydrocarbons will be converted into naphtha boiling range paraffins for subsequent aromatic production via dehydrocyclization.

When operating a catalytic reforming system at a relatively high severity level, the paraffinic hydrocarbons are subjected to hydrocracking. Although this reduction in the molecular weight of the paraffinic hydrocarbons partially increases the oxtane rating of the normally liquid gasoline boiling range products, substantial quantites of gaseous material are produced. This light gaseous material is substantially completely saturated and comprises methane, ethane, propane and butane. As hereinafter set forth, the propane/butane concentrate may be recovered and utilized in an alkylation reaction system to produce a normally liquid, highoctane alkylate product. At a relatively low reforming severity, paraffin cracking is decreased with the result that an increased quantity of low octane rating saturates is produced. In order to upgrade the overall quality of the gasoline pool, either the addition of lead becomes necessary, or the low octane rating staturates must be subjected to further processing to produce higher octane components. This type of operation produces a two-fold effect, notwithstanding an increase in the octane rating of the final product; first, additional high-octane aromatic components are produced and, secondly, the low-octane rating components are at least partially eliminated by conversion either to aromatic hydrocarbons, or to light normally gaseous material. The inclusion of the higher molecular weight, low octane rating paraffins requires that more aromatics be produced to achieve the desired octane rating of the final product, and thus creates a larger liquid yield loss. The results include a lower liquid volumetric yield due both to shrinkage in molecular size when paraffins and naphthenes are converted to aromatics, and to the production of the aforesaid light gaseous components. In accordance with one overall refinery operation, into which the present invention is integrated, the hydrocracking system is dove-tailed with at least a separation system, a catalytic reforming unit and an alkylation unit. The end result is the production of a high-octane, unleaded gasoline pool, unaccompanied by substantial liquid yield loss.

I-Iydrocarbonaceous charge stocks, contemplated for conversion in accordance with the present invention, contain kerosene boiling range constituents. In the petroleum refining art, gasoline boiling range hydrocarbons generally connotes those hydrocarbons having an initial boiling point of at least about 100 F., being inclusive of intermediate boiling range fractions referred to as light naphtha and heavy naphtha. Contemplated charge stocks, for use in the present invention will have an end boiling point above 410 F., often being as high as about 600 F., although a maximum of about 550 F. would be preferred. It is not intended to limit the present invention to a charge stock having a particular boiling range. Suffice to say, a suitable charge stock will generally have an initial boiling point about about F. and an end boiling point below about 600 F. A particularly preferred charge stock would be a full boiling range naphtha in admixture with a kerosene cut. That is, a charge stock containing C hydrocarbons and having an end boiling point of about 550 F. to 600 F.

In those instances where the reformed product effluent is subjected to further separation to recover a substantially pure aromatic concentrate, the paraffinic hydrocarbons may be recycled to the hydrocracking reaction zone. In this manner, the charge stock is reduced in boiling range, while the ring structure and eventual aromatic yield are maintained. The precise boiling range of any given charge stock will be primarily dependent upon the economic and processing considerations prevalent in the particular locale where the charge stock is available.

OBJECTS AND EMBODIMENTS A principal object of the present invention constitutes the simultaneous production of aromatic hydrocarbons and an isobutane concentrate.

Another objective resides in the production of a highoctane, unleaded motor fuel gasoline pool. A corollary objective of our invention is to provide an integrated refinery operation for producing high liquid yields of a high-octane, unleaded gasoline pool from hydrocarbonaceous material boiling above about 400 F.

Therefore, in one embodiment, the present invention involves a process for the simultaneous production of an aromatic concentrate and isobutane, from a charge stock containing hydrocarbons boiling above a temperature of about 400 F., which process comprises the steps of: (a) separating said charge stock, in a first separation zone, to provide a first fraction having an end boiling point in the range of about 225 F. to about 260 F., a second fraction having an initial boiling point in the range of 225 F. to 260 F. and an end boiling point in the range of about 300 F. to about 410 F., and a third fraction having an initial boiling point in the range of about 300 F. to about 410 F. and an end boiling point greater than about 410 F.; (b) reacting said first and third fractions with hydrogen, in a hydrocracking reaction zone, at hydrocracking conditions and in contact with a hydrocracking catalytic composite of a Group VIII noble metal component, and the reaction product of alumina and a gem-polyhalo compound; (c) separating the resulting hydrocracked product effluent, in a second separation zone, to provide a heptane-plus concentrate and to recover said isobutane; (d) reacting said heptane-plus concentrate and said second fraction in a catalytic reforming zone, at reforming conditions selected to convert naphthenic and paraffinic hydrocarbons to aromatic hydrocarbons; and, (e) recovering said aromatic concentrate from the resulting reformed product effluent.

Other embodiments of our invention involve the use of various catalytic composites, operating conditions and processing techniques. In one such other embodiment, the reformed product effluent is separated, in a third separation zone to provide a saturated paraffinic stream and to recover the aromatic concentaate. In another such embodiment, the paraffinic stream, recovered from the third separation zone, may be reacted with hydrogen in said hydrocracking reaction zone.

In a more limited embodiment, a pentane/hexane concentrate is separated from the hydrocracked prodnet effluent and subsequently reacted with hydrogen in a hydroisomerization reaction zone to produce pentane and hexane isomers. Similarly, a (l /C concentrate may be recovered from the reformed product effluent and subjected to hydroisomerization. In still another embodiment, at least a portion of isobutane concen trate is reacted with an olefinic hydrocarbon, in an alkylation reaction zone, to produce a normally liquid alkylate product.

In a preferred embodiment, the end boiling point of the second fraction is in the range of about 380 F. to 410 F., the initial boiling point of the third fraction is about 380 F. to 410 F., while its end boiling point can be as high as 550 F. to about 600 F.

SUMMARY OF INVENTION As hereinbefore set forth, the present invention involves a particular saturate cracking zone and a catalytic reforming zone. Additionally, in another embodiment, an integrated refinery scheme, incorporating the present inventive concept, utilizes a solvent extraction zone, an isomerization zone and an alkylation reaction zone. In a specific embodiment, the overall refinery process includes a dehydrogenation reaction zone to produce the olefmic hydrocarbons utilized in the alkylation reaction zone. In order that a clear understanding of the integrated refinery process be obtained, a brief description of each of the various reaction and separation zones, utilized in one or more embodiments, is believed to be warranted. In describing each individual zone, one or more references will be made to United States Patents in order that more details will be available where desired. Such references are not to be construed as either exhaustive, or limiting, but merely exemplary and illustrative.

We have observed that, notwithstanding the preservation of the ring compounds in the hydrocracking reaction zone, alkyl groups will be removed from those ring compounds containing the same. These alkyl groups contribute to the net liquid yield loss in the form of light gaseous material; however, a more significant consideration is that the aromatic hydrocarbons in the ultimate product will contain fewer alkyl groups and possess, therefore, a lower clear octane blending value. For example, benzene has a clear blending value of about 99, ethylbenzene a value of about I2 while the various xylenes have values ranging from about R to about 150. These clear octane blending values, as well as those of other hydrocarbon species, are obtaiend from API Research Project Number $5, Tabulated Knock-test Data to June 30, 1954, Table IV, page I 19. It is noted, therefore, that the removal of the alkyl groups from the ring compounds results in a lower clear octane blending value.

In accordance with the present invention, the charge stock is initially separated, for example in a fractionation system, to provide three fractions having the boiling range characteristics hereinbefore set forth. The precise split among the three fractions will be principally dependent upon the chemical characteristics of the feed stock, the results of its detailed component analysis, and the desired product specifications, particularly octane rating and end boiling point. The first and third fractions serve as the feed stock to the hydrocracking reaction zone. Through the utilization of the present invention, 400 F.-plus material can be reduced in molecular weight, and thus in boiling range, while maintaining the desired cyclic structure for eventual aromatic production. A possible lower end point specification for motor fuel, in the future, can thus be easily met while simultaneously distributing the high octane values more thoroughly throughout the final liquid product.

In a preferred mode of operation, the separation is effected such that substantially all of the C -ring compounds in the fresh feed charge stock are withdrawn with the second fraction and introduced into the catalytic reforming reaction zone. In those instances where the catalytically reformed product efiluent is subjected to a solvent extraction technique to recover the aromatic concentrate, the paraffinic raffinate may be recycled to the hydrocracking reaction zone, and/or to the isomerization zone, for conversion therein into valuable normally liquid isomeric hydrocarbons.

HYDROCRACKING REACTION ZONE The charge stock processed in accordance with the present invention, may be derived from a multitude of sources. For example, one source constitutes those distillates obtained from full boiling range petroleum crude oils; another source is the fraction obtained from the catalytic cracking of gas oils and other heavier hydrocarbon mixtures, while another source constitues the C 600 F. boiling range effluent from a hydrocracking reaction zone which processes heavier, gas oil feed stocks. In view of the fact that the greater proportion of such fractions are contaminated through the in clusion of sulfurous and nitrogenous compounds, it is contemplated that the overall process may have integrated therein a hydrorefining zone, details of which are well known and thoroughly described within the prior art. It is understood that such hydrogenative pretreatment of the charge stock does not constitute an essential feature of the present combination process, but is a preferred technique.

As hereinbefore set forth, one key feature of the present inventive concept resides in the use of a particular hydrocracking reaction zone. This hydrocracking reaction zone is unlike those in current use, both in function and end result. Initially, the charge to the hydrocracking reaction zone constitutes the lower and higher boiling portions of the charge stock, and the product effluent contains very little, if any, methane, ethane and propane. Any propane in the product effluent can be recovered and subsequently utilized as LPG, or as feed to the alkylation reaction zone, or for the synthesis of isopropyl alcohol. It is understood that the hydrocracking reaction zone may consist of separate reactor vessels whereby the first and third fractions can be individually reacted with hydrogen. Similarly, the reaction zone may be a single vessel having multiple catalyst beds disposed therein. Through the utilization of a particular catalytic composite and operating conditions, the cracking of paraffinic hydrocarbons in the charge stock produces relatively large quantities of butane, exceedingly rich in isobutane. Similarly, virgin pentanes and hexanes are converted to various isomers thereof. With respect to the overall octane rating of the gasoline pool, this aspect is of great significance. Whereas n-pentane has an octane blending value of only 62, isomers thereof average about 95.9; similarly, although n-hexane has a clear octane blending value of about 20, the isomers thereof have an average rating of about 90.8.

ln view of the unique character of the product effluent, being exceedingly rich in isoparaffins, the present hydrocracking reaction zone is herein referred to as lcracking. The selective nature of the hydrocracking reactions taking place include the retention of cyclic rings and a reduction in molecular weight of those rings via isomerization and the splitting of isobutane from the parent molecule. Thus, polycyclic compounds boiling in the higher temperature range of the feed stock are converted to lower-boiling naphthenes which are, in turn, converted into gasoline boiling range aromatics in the subsequent catalytic reforming reaction zone. Benefits are thus afforded since the high octane aromatic hydrocarbons will be more uniformly distributed throughout the final gasoline boiling range product.

The hdyrocracking reaction conditions, under which the process is conducted, will vary according to the chemical and physical characteristics of the particular charge stock. Hydrocracking reactions have heretofore generally been effected at pressures in the range of about 1,500 to about 5,000 psig., a liquid hourly space velocity of about 0.25 to about 5.0, hydrogen circulation rates of about 5,000 to about 50,000 scf./Bbl. and maximum catalyst bed temperatures in the range of about 700 F. to about 950 F. As discussed in the prior art, heavier charge stocks require a relatively high severity of operation including high pressures high catalyst bed temperatures and realtively low liquid hourly space velocities. A lower severity of operation is employed with comparatively lighter feed stocks such as the kerosenes and light gas oils. ln accordance with the present invention, regardless of the precise characteristics of the charge stock, the hydrocracking process is effected at a relatively lower severity than is now commonly in use. The hydrocracking reaction zone has disposed therein a catalytic composite comprising the reaction product of a Group Vlll noble metal component and alumina with a gem-polyhalo compound. The conversion conditions include a liquid hourly space velocity of about 0.5 to about 10.0, a hydrogen circulation rate of about 3,000 to about 20,000 scf./Bbl., a pressure from about 200 to about 2,000 psig., and preferably less than about 1,000 psig. and, of greater significance, a maximum catalyst bed temperature from about 300 F. to about 600 F. ln many instances, the operating pressure will consistently be in the range of about 200 to about 500 psig., the hydrogen concentration is from about 3,000 to about 10,000 scf./Bbl. and the liquid hourly space velocity about 2.0 to about 10.0 without inducing serious effects either in regard to the effective life of the catalytic composite, or with respect to the desired product slate.

As hereinbefore set forth, the hydrocracking reaction zone has disposed therein a catalytic composite containing the reaction product of a Group Vlll noble metal component and alumina with a gem-polyhalo compound; the latter may be a sublimed Friedel-Crafts metal halide. The reaction of the gem-polyhalo compound with the alumina/Group Vlll noble metal composite produces catalytic sites of very high acidity. When the gem-polyhalo compound is aluminum chloride, there exists some evidence that the active sites are characterized by the following grouping:

LII

These react reversibly with hydrochloric acid, generally added during processing, to produce strong protonic sites as:

10 ble metal composite may be treated with more than one gem-polyhalo compound, such treatment being effected either simultaneously, or sequentially. In view of the fact that the carrier material originally contains alumina, the mechanism of the reactions with the gem- 15 polyhalo compound, which is not now known with ac- -curacy, produces a type of group as above indicated.

Considering first the porous carrier material, it is preferred that it be adsorptive and passes a high surface area. Heretofore, carrier materials have been selected from the group of amorphous refractory inorganic oxides including alumina, titania, zirconia, silica, mixtures thereof, etc. When of the-amorphous type, the preferred carrier material appears to be a composite of alumina and silica, with the latter being present in an amount of about 10.0 percent to about 90.0 percent by weight. Recent developments in the area of catalysis have further shown that various crystalline aluminosilicates can be employed to advantage in some hydrocracking situations. Such zeolitic material includes mordenite, faujasite, Type A or Type U molecular sieves, etc. In view of the fact that a gem-polyhalo compound does not appear sufficiently strong to react with silica, the preferred carrier material constitutes alumina. While the action and effect on refractory material other than alumina and silica, for example zirconia, is not known with accuracy, it is believed that reaction does not take place to a degree sufficient to produce the desired catalyst and result.

The hydrocracking catalytic composite contains a Group Vlll noble metal component. Suitable metals are those from the group including platinum, palladium, rhodium, ruthenium, osmium, and iridium, with platinum and palladium being particularly preferred. These metal components, for example platinum, may exist within the final composite as a compound such as an oxide, sulfide, halide, etc., or in an elemental state, the latter being preferred, Generally, the amount of the noble metal component is small compared to the quantities of the other components combined therewith. On an elemental basis, the noble metal component comprises from about 0.01 percent to 2.0 percent by weight of the final catalytic composite. Bi-metallic catalysts, containing germanium, rhenium, or tin, in addition to the noble metal, are also suitable for use in the hydrocracking reaction zone.

The catalytically active metallic components may be incorporated within the catalytic composite in any suitable manner including co-precipitation or cogellation with the carrier material, ion-exchange or impregnation. The latter constitutes a preferred method of preparation, and utilizes water-soluble compounds of the metallic components. Thus, a platinum component may be added to a carrier material by commingling the latter with an aqueous solution of chloroplatinic acid. Other water-soluble compounds may be employed, and include ammonium chloroplatinate, platinum chloride, chloropalladic acid, palladic chloride, etc. Following impregnation, the carrier material is dried and subjected to a calcination, or oxidation technique which is followed by reduction in hydrogen at some elevated temperature, prior to contact by the selected gempolyhalo compound.

The method of incorporating a Freidel-Crafts metal halide, as the gem-polyhalo compound, involves a sublimation, or vaporization technique, with the vaporized metal halide contacting the alumina containing the Group VIll noble metal component. That is, the catalytically active metallic component is composited with the alumina and reduced, prior to contact with the gem-polyhalo compound. Briefly, therefore, the preferred technique involves the treatment with a gempolyhalo compound after the catalytically active metal components have been impregnated onto the carrier material, and following drying, calcination and reduction in hydrogen. Where the gem-polyhalo compound is a Friedel-Crafts metal halide, such as aluminum chloride, it is vaporized onto the carrier and heated to a temperature of about 300 C. for a time sufficient to remove any unreacted metal halide. Thus, the final catalytic composite does not contain any free, uncombined Friedel-Crafts metal halide. Following vaporization of the metal halide and heating of the thus-formed composite, the refractory inorganic oxide will be increased in weight by from about 2.0 percent to about 25.0 percent, based upon the original weight of the carrier material. While the exact increase in weight does not appear to be critical, high activity catalysts are obtained when the thus-treated refractory material has a weight increase of about 5.0 percent to about 20.0 percent. Further details of this sublimation technique may be found in U. S. Pat. No. 2,924,628 (Class 260-666).

Various Friedel-Crafts metal halides may be utilized as the gem-polyhalo compound, but not necessarily with equivalent results. Examples of such metal halides include aluminum bromide, aluminum chloride, antimony pentachloride, beryllium chloride, germanium tetrachloride, ferric chloride, ferric bromide, gallium trichloride, stannic bromide, stannic chloride, tetanium tetrabromide, tetanium tetrachloride, zinc bromide, zinc chloride, and zirconium chloride. The Friedel- Crafts aluminum halides are preferred, with aluminum chloride and/or aluminum fluoride being particularly preferred. This is due to the ease of preparation and the fact that the thus-prepared catalysts have an unexpectedly high activity for the selective production of isoparaffms, and particularly for isobutane.

Temperatures at which the Friedel-Crafts metal halide is vaporized onto the alumina will vary in accordance with the particular metal halide utilized. Vaporization is carried out either at the boiling, or sublimation point of the particular friedel-crafts metal halide, or at some temperature not substantially exceeding these points; for example, not more than about 100 C. higher than the boiling point or sublimation point. Aluminum chloride sublimes at a temperature of about 178 C. and a suitable vaporization temperature, utilizing this gem-polyhalo compound, will, therefore, range from about l80 C. to about 275 C.

As hereinbcfore stated, the catalyst may comprise the reaction product of an alumina-Group Vlll noble metal composite and one or more gem-dihalo or gempolyhalo compounds. The interaction produces catalytic sites of higher acidity than can be produced, for example, by treatment with hydrochloric acid. Such compounds include carbon tetrachloride, CHCl sulfur dichloride, sulfur oxychloride, PCl POCl R-CH-Cl etc. Therefore, the catalyst will possess the desired properties to produce isomeric parafiins while simultaneously reducing the molecular weight of naphthenic compounds. Prior to its use, the catalytic composite is subjected to a substantially water-free reduction technique. This is designed to insure a more uniform and thorough destribution of the metallic components throughout the carrier material. Substantially pure and dry hydrogen is employed as the reducing agent at a temperature of about 800 F. to about l,200 F., and for a time sufficient to reduce the metallic components.

In view of the fact that the reactions being effected are exothermic in nature, an increasing temperature gradient is experienced as the reactants traverse the catalyst bed. In accordance with the present process, the maximum catalyst bed temperature is maintained in the range of about 300 F. to about 600 F. In order to assure that the maximum catalyst bed temperature does not exceed the allowable limit, the use of conventional quench streams, either normally liquid, or normally gaseous, introduced at one or more intermediate loci of the catalyst bed, is contemplated.

As hereinbefore set forth, the product effluent from the hydrocracking reaction zone, considering only the normally gaseous portion thereof, is predominatnly butane, the greater proportion of which constitutes isobutane. Similarly, the pentane/hexane concentrate is rich in isomers of higher octane rating; for this reason, the hydrocracking reaction zone is herein referred to as lcracking, the l alluding to isomer production. In addition to the production of exceedingly large quantities of isobutane, accompanied by little yield loss to methane and ethane, an unusual and unexpected result is the virtually complete retention of cyclic hydrocarbons originally present in the fresh feed charge stock. Furthermore, those heavier cyclic hydrocarbons in the fresh feed have been reduced in molecular weight such that the subsequent reformed product effluent exhibits a more uniform distribution of the high octane aromatic components. This achieves importance from the standpoint of the possible lowering of the end boiling point of motor fuel at some future date.

The foregoing is evidenced by results which were obtained when a Mid-continent, straight-run naphtha fraction was subjected to l-cracking. This charge stock had a gravity of 55.0 API, an initial boiling point of about 2l0 F. and an end boiling point of about 269 F. Virgin cyclic hydrocarbons in the charge stock constituted 52.6 percent by weight. The l-cracking yields, being based on weight percent of the naphtha charge stock were as follows: butanes-minus, 21.2 percent by weight; pentanes, 11.7 percent; hexanes, l 1.9 percent; and, heptane-plus hydrocarbons, 56.5 percent, which values are inclusive of a hydrogen consumption in the amount of about 1.3 percent by weight of the naphtha charge stock. From the standpoint of selectivity, butanes were produced in an amount of 92.0 percent by weight of the total butane-minus portion, the isobutane content of the total butanes being 92.0 percent by weight; of the total pentanes produced, 89.0 percent by weight were isomeric in nature. On a molar basis, the cyclic retention amounted to 99.0 percent.

As previously set forth, the selective l-cracking operation also has an effect on the boiling range of the aromatics produced in the subsequent catalytic reforming step. When a comparison is made between the product produced by direct catalytic reforming of the Midcontinent charge stock, and the product resulting from I-cracking the naphtha followed by catalytically reforming the heptane-plus portion of the hydrocracked product, it is noted that the last 50.0 percent by volume indicates a lower boiling range to the extent that there is a 40 F. difference in the end boiling point; the end boiling point is, in fact, lower than that of the original feed stock. Catalytic reforming of itself results in a product having a gravity of 43.2 AP1, an initial boiling point of 144 F. and an end boiling point of 404 F., having a clear octane rating of 96.3 and containing 61.8 percent by volume of aromatic hydrocarbons. Where the naphtha charge stock is initially subjected to l-cracking, followed by the catalyst reforming (at the same reforming conditions) of the heptane-plus portion, the product has a gravity, AP1 of 36.5, an initial boiling point of 162 F. and an end boiling point of 358 F., possesses a clear octane rating of 105.4 and contains 84.5 percent by volume of aromatic hydrocarbons. In this particular comparison, the lower boiling front end of the product resulting from the combination cannot be compared directly since it does not include the pentane/hexane portion of the hydrocracked product effluent. The overall yields, considering a butane-plus product in an amount of 101.2 percent by volume, based upon the naphtha charge stock, having a clear research octane rating of 98.9, is produced with an accompanying yield loss to methane, ethane and propane of about 4.0 percent by weight.

One of the principal objects of the present inventive concept is to afford a method for achieving a distinct improvement in the foregoing results while including high-boiling components in the feed stock. Although the degree of cyclic hydrocarbon retention is approximately the same, and the end product consists of a similar volumetric quantity of aromatic hydrocarbons, the octane blending values are increased as a result of a greater degeree of retention of alkyl groups on the alkyl-substituted ring compounds. Furthermore, there exists a lower liquid yield loss resulting principally from a decerased production of metahne and ethane.

CATALYTIC REFORMING ZONE The product effluent from the l-cracking reaction zone is separated, in suitable fractionation facilities, into various component streams. A butane concentrate, consisting predominantly is isobutane, is recovered and subjected to either alkylation, or dehydrogenation as hereinafter set forth. A pentane/hexane concentrate, rich in isomers thereof, is separately recovered and may be introduced directly into the unleaded gasoline pool. In one embodiment, normal pentane and normal hexane are separately recovered and subjected to isomerization to produce additional pentane/hexane isomers. The heptane-plus portion of the l-cracked product effluent constitutes a portion of the charge to the catalytic reforming reaction zone, being commingled with the second fraction obtained from the first separation zone.

Catalytic composites, suitable for utilization in the reforming reaction zone, generally comprise a refractory inorganic oxide carrier material containing a metallic component selected from the noble metals of Group VIII. Recent developments have indicated that activity and stability are significantly enhanced through the addition of various catalytic modifiers, especially tin rhenium, nickel and/or germanium, thereby forming multi-metallic catalysts. Suitable porous carrier materials include refractory inorganic oxides such as alumina, silica, zirconia, etc. Generally favored metallic components include ruthenium, rhodium, palladium, osmium, rhenium, platinum, iridium, germanium, nickel and tin, and mixtures thereof. These metallic components are employed in concentrations ranging from about 0.01 percent to about 5.0 percent by weight, and preferably from about 0.01 percent to about 2.0 percent by weight. Reforming catalysts may also contain combined halogen selected from the group of chlorine, fluorine, bromine, iodine and mixtures thereof, with chlorine and fluorine being particularly preferred.

Since the reforming reaction zone processes substantially only a heptane-plus product, the conditions required result in a lower degree of operating severity. Operating severity level increases with increased temperature and lower liquid hourly space velocity. Those skilled in the reforming art will immediately recognize that a lower severity operation is more advantageous from the standpoint of catalyst life and decrease production of methane and ethane. Furthermore, at a higher space velocity, the refiner enjoys the added advantage of greater throughput per unit of time.

Illustrative catalytic reforming processes are found in U. S. Pat. Nos. 2,095,620 (Class 208-65), 3,000,812 (Class 208-138) and 3,296,118 (Class 208-). Effective prior art reforming operating conditions include a catalyst temperature within the range of about 850 F. to about 1,100 F., a liquid hourly space velocity of about 1.0 to about 5.0 and a pressure of about 500 to about 1,000 psig. The quantity of hydrogen-rich gas, in admixture with the hydrocarbon charge stock, is generally in the range of about 1.0 to about 20.0 moles of hydrogen per mole of hydrocarbon. In accordance with the present combination process, the catalyst refonning reaction zone will normally function at a temperature in the range of about 800 F. to about 1,000 F., a liquid hourly space velocity of abou2.0 to about 10.0, or more, and a pressure in the range of about 100 to about 500 psig. The hydrogen concentration will generally be in the same range as that of the prior art. The product effluent from the catalytic reforming reaction zone is generally introduced into a high-pressure separation zone at a temperature in the range of about 60 F. to about 140 F. to separate lighter components from heavier, normally liquid components. Since normal reforming operations are hydrogen-producing, a ceratain amount of a hydrogen-rich stream is generally removed from the reforming system by way of pressure control. It is within the scope of the present invention that such excess hydrogen be employed in the hydrogenconsuming hydrocracking reaction zone, as make-up hydrogen, as well as in the hydroisomerization reaction zone.

AROMATIC SEPARATION ZONE As hereinabove set forth, the reforming reaction zone is maintained at relatively low severity operating conditions in order to produce a product effluent rich in aromatic hydrocarbons, and for the purpose of dehydrocyclization of the paraffinic material in the charge stock. Although any aromatic separation scheme such as fractionation, may be utilized, a greater degree of efficiency is achieved through the use of a solvent extraction system. Solvent extraction, to produce an aromatic concentrate and a paraffinic raffinate, is a well known technique thoroughly described in the literature. Suitable schemes involve the operations illustrated in U. S. Pat. Nos. 2,730,558 (Class 260-674) and 3,361,664 (Class 208-313). Solvent extraction processes utilize a solvent which possesses a greater selectivity and solvency for aromatic components in the reformed product effluent than for the paraffinic components. Selective solvents may be selected from a wide variety of normally liquid organic compounds of generally polar character; that is, compounds containing a polar radical. In any given situation, the particular solvent is one which boils at a temperature above the boiling point of the hydrocarbon mixture at an ambient extraction pressure. Illustrative specific organic compounds, useful as selective solvents in extraction processes for the recovery of aromatic hydrocarbons, include the alcohols, such as glycols, including ehtylene glycol, propylene glycol, butylene glycol, tetra-ehtylene glycol, glycerol, diethylene glycol, dipropylene glycol, tripropylene glycol, etc.; other organic solvents well known in the art, for extraction ofhydrocarbon components from mixtures thereof with other hydrocarbons, may be employed. A particularly preferred class of such other solvents are those characterized as the sulfolane-type. Thus, as indicated in U.S. Pat. No. 3,470,087 (Class 208321), a preferred solvent is one having a fivemembered ring, one atom of which is sulfur, the other four being carbon and having two oxygen atoms bonded to the sulfur atom. In addition to sulfolane, the preferred class includes the sulfolenes, such as 2- sulfolene and 3-sulfolene.

Selectivity of the foregoing described solvents may be enhanced further through the addition of water. This increases the selectivity of the solvent phase for aromatic hydrocarbons over non-aromatic hydrocarbons. As a general practice, the solvent composition will contain from about 0.5 percent to about 20.0 percent by weight of water, and preferably from about 2.0 percent to about 15.0 percent by weight, principally depending on the particular solvent and the process conditions under which the extraction, extractive distillation and solvent recovery zones are functioning.

in general, solvent extraction is effected at elevated temperatures and pressures which are selected to maintain the charge stock and solvent in a liquid phase. Suitable temperatures are within the range of about 80 F. to about 400 F., and preferably from about 150 F. to about 300 F. Operating pressures include superatmospheric pressures up to about 400 psig. and preferably from about 15.0 psig. to about 150 psig. Extractive distillation zone pressures are from atmospheric to about 100 psig., although the pressure at the top of the distillation zone will generally be maintained in the range of about 1.0 psig. to about psig. The reboiler temperature is dependent upon the composition of the feed stock and the selected solvents, although temperatures from about 275 F. to about 360 F. appear to yield more satisfactory results. The solvent recovery system is operated at relatively low pressures and sufficiently high temperatures to drive the aromatic hydrocarbons overhead, thus producing a lean solvent bottoms stream. Preferably, the top of the solvent recovery zone is maintained at a pressure from about 100 to about 400 mm. Hg., absolute. These low pressures must be utilized since the reboiler temperatures should be maintained below about 370 F. in order to avoid thermal decomposition of the organic solvent.

ALKYLATlON REACTION ZONE Since the preferred use of the present inventive concept constitutes the integration thereof into an overall refinery scheme for the production of voluminous quantities of a high-octane, unleaded motor fuel gasoline pool, the isobutane-rich, butane concentrate from the l-cracking zone is utilized as fresh feed charge stock to an alkylation reaction zone. Alkylation is effected by intimately commingling the isobutane feed, an olefinic hydrocarbon and a particular catalyst as hereinafter described. It is understood that the source of the olefinic hydrocarbon, for utilization in the alkylation reaction zone, is not essential to the process encompassed by the present invention. Thus, outside olefinic material may be supplied from any suitable source such as a fluid catalytic cracking unit, or a thermal cracking unit. However, as indicated in another specific embodiment, at least a portion of the isobutane concentrate may be subjected to dehydrogenation in a dehydrogenation reaction zone to produce the alkylatable olefinic hydrocarbons. In still another embodiment, the propane produced within the process may also be dehydrogenated and introduced into the alkylation reaction zone.

The alkylation reaction zone may be any acidic catalyst reaction system such as a hydrogen fluoride, or sulfuric acid system, or one which utilizes a boron halide in a fixed-bed reaction system. Hydrogen fluoride alkylation is particularly preferred, and may be conducted substantially as set forth in US. Pat. No. 3,249,650 (Class 260-68348). Briefly, the alkylation conditions, when effected in the presence of hydrogen fluoride catalyst, are such that the catalyst to hydrocarbon volume ratio within the alkylation reaction zone is in the range of about 0.5 to about 2.5. Ordinarily, anhydrous hydrogen fluoride will be charged to the alkylation system as fresh catalyst; however, it is possible to utilize hydrogen fluoride containing as much as about 10.0 percent'water, although excessive dilution with water is generally avoided since it tends to reduce the alkylating activity of the catalyst and further introduces a variety of corrosion problems into the process. In order to reduce the prevailing disposition of the olefinic portion of the charge stock to undergo egregious polymerization prior to alkylation, the molar proportion of isoparaffin to olefinic hydrocarbon with the alkylation reaction zone is desirably maintained at a value greater than 1.0, and preferably from about 3.0 to about 15.0. Alkylation reaction conditions, when catalyzed by hydrogen fluoride, include a temperature from 0 to about 150 F. and preferably from about 30 F. to about F. The pressure maintained within the alkylation system is ordinarily at a level sufficient to maintain the hydrocarbons and catalyst in substantially liquid phase; that is, from about atmospheric to about 40 atmospheres. The contact time within the alkylation reaction zone is conveniently expressed in terms of space-time, and is generally defined as the volume of catalyst within the contact zone divided by the volume rate per minute of hydrocarbon reactants charged to the zone. Usually, the space-time factor will be less than 30 minutes and preferably less than about 15 minutes.

The alkylation reaction zone effluent is separated to provide an acid phase and a hydrocarbon phase, the

latter being separated to recover the normally liquid alkylate product and unreacted isobutane. The alkylate product, in combination with the aromatic concentrate from the solvent extraction zone forms part of the unleaded gasoline pool. Unreacted isobutane may be recycled to the alkylation reaction zone, or a portion thereof may be diverted to the dehydrogenation reaction zone for the purpose of producing additional olefinic hydrocarbons for utilization in the alkylation reaction zone.

ISOMERIZATION REACTION ZONE The n-pentane and n-hexane separated from the product effluent of the I-cracking reaction zone and/or the reforming zone, possess clear research blending values of about 62 and respectively. These components are not, therefore, desirable in a gasoline pool which is intended to be free from lead additives. Therefore, in still another embodiment of the present invention, a normal pentane/normal hexane stream is separately recovered and introduced into an isomerization reaction zone for the purpose of producing an effluent product rich in pentane and hexane isomers. Since the retention of carbon numbers (as charged) in the isomerization reaction zone is virtually 100.0 percent, the unleaded gasoline pool can be significantly increased in its clear research octane rating without incurring a detrimental volumetric yield loss.

As indicated in U.S. Pat. No. 3,131,235 (Class 260-6833), the isomerization process is effected in a fixed-bed system utilizing a catalytic composite of a refractory inorganic oxide carrier material, a Group VIII noble metal component and combined halogen, preferably selected from fluorine and chlorine. The refractory inorganic oxide carrier material may be selected from the group including alumina, silica, titania, zirconia, mixtures of two or more, and various naturallyoccurring refractory material. Of these, a syntheticallyprepared, gamma-alumina is preferred. The Group VIII noble metal is generally present in an amount of about 0.01 percent to about 2.0 percent by weight, and may be one or more metals selected from the group of ruthenium, rhodium, osmium, iridium, and particularly platinum and/or palladium. The amount of combined halogen will be varied from about 0.01 percent to about 8.0 percent by weight. Both fluorine and chlorine may be used to supply the combined halogen, although the use only of fluorine in an amount of about 2.5 percent to about 5.0 percent by weight is preferred.

The isomerization reactions are preferably effected in a hydrogen atmosphere utilizing sufficient hydrogen so that the hydrogen to hydrocarbon mole ratio to the reaction zone will be within the range of about 0.25 to about 10.0. Operating conditions will additionally include temperatures ranging from about 200 F. to about 600 F., although temperatures within the more limited range of about 230 F. to about 320 F. will generally be employed. The pressure, under which the reaction zone is maintained, will range from about 400 to about 1,000 psig., and the liquid hourly space velocity from 1.0 to about 3.0. Hydrogen is separated from the reaction products and recycled, while the normally liquid effluent is subjected to fractionation and separation to produce the desired isomerized product. Recovered starting material is also recycled to the reaction zone to increase the overall process yield. Another suitable isomerization process is that described in U.S. Pat. No. 2,924,628 (Class 260-666).

DEHYDROGENATION REACTION ZONE At least a portion of the recovered butane concentrate may be subjected to dehydrogenation to produce the olefin required for alkylation within the alkylation reaction zone. In another embodiment, the propane stream is also subjected to dehydrogenation to provide additional olefins. The advisability of the utilization of either, or both techniques will be primarily dependent upon the availability of outside olefins; for example, from a catalytic or thermal cracking unit.

When dehydrogenation is deemed desirable, it may be effected essentially as set forth in U.S. Pat. No. 3,293,219 (Class 260-683). Briefly, dehydrogenation reactions are generally effected at conditions including a temperature in the range of from 400 C. to about 700 C., a pressure from about atmospheric to about psig., a liquid hourly space velocity within the range of about 1.0 to about 40.0 and in the presence of hydrogen in an amount to result in a mole ratio of from l.0:l0.0 to l0.0:l.0, based upon the paraff'mic charge.

The dehydrogenation catalyst is a composite of an inorganic oxide carrier material, an alkali metal component, a Group VIII metal component and a catalytic modifier from the group consisting of arsenic, antimony, bismuth, rhenium, germanium and tin. A preferred catalyst comprises Iithiated alumina containing from 0.05 percent to about 5.0 percent by weight of a Group VIII noble metal, particularly platinum. The catalytic modifier is employed in an amount based upon the concentration of Group VIII noble metal component. For example, arsenic is present in an atomic ratio of arsenic to platinum in the range of about 0.20 to about 0.45. Although lithium is the preferred alkalinous metal component, the catalyst may contain calcium, magnesium, strontium, cesium, rubidium, potassium, sodium and mixtures thereof, etc.

Dehydrogenation conditions and catalysts result in a relatively low equilibrium conversion per pass, accompanied by relatively high selectivity to the desired olefinic hydrocarbons. Thus, while the conversion per pass might range from about 10.0 percent to about 35.0 percent, the selectivity of conversion will range from about 93.0 percent to about 97.0 percent, or higher. In view of the fact that the alkylation reactions are effected with a molar excess of paraffins over olefinic hydrocarbons, the high selectivity and relatively low conversion in the dehydrogenation zone are advantageous.

DESCRIPTION OF DRAWING The inventive concept, encompassed by the present process, and one embodiment are illustrated and described in the accompanying drawing. The illustration is presented by way of a simplified block-type flow diagram representing the I-cracking and catalytic reforming reaction zones. Also illustrated are two fractionating columns utilized in the separation of the fresh feed charge stock and the product effluent from the I- cracking reaction zone. Miscellaneous appurtenances, not believed necessary for a completely clear understanding of the present process, have been eliminated. The utilization of details such as pumps, compressors, instrumentation and controls, heat-recovery circuits, miscellaneous valving, start-up lines and similar hardware, etc., is well within the purview of those skilled in the petroleum refining art. Similarly, with respect to the flow of materials throughout the system, only those major streams required to illustrate the interconnection and interaction of the various zones are presented. Thus, various recycle lines and vent-gas streams, for example, have also been eliminated.

With reference now to the drawing, it will be described by way of processing a naphtha/kerosene fraction obtained from a Mid-continent vacuum gas oil. This charge stock, C 550 F., is initially subjected to hydrogenative hydrorefining for nitrogen and sulfur removal. The charge stock is introduced into the process by way of line l, and is separated, in fractionator 2, to provide a heart-cut naphtha fraction (230 F.) in line 4 and a C 230 F. light fraction in line 5. The light and heavy fractions are commingled and continue through line 5 into l-cracking reaction zone 6. The catalytic composite, disposed in ll-cracking zone 6, constitutes alumina, 0.75 percent by weight of platinum and 7.5 percent by weight of aluminum chloride, sublimed thereon to react with the alumina. Operating conditions include a pressure of about 450 psig. a maximum catalyst bed temperature of about 425 F., a liquid hourly space velocity of about 1.5 and a hydrogen to hydrocarbon mole ratio of about 60:10.

The product effluent is withdrawn from I-cracking zone 6 by way of line '7, and introduced therethrough into fractionator 0. Butanes and lighter components are withdrawn through line 9 while a pentane/hexane concentrate is withdrawn by way of line 10. A C 390 P. fraction is removed by way of line lll, combined with the heavy naphtha fraction in line 3 (from fractionator 2), the mixture continuing through line 3 into catalytic reforming zone 12. Kerosene boiling range material is withdrawn as a bottoms stream from fractionator 0 through line 113, and recycled to l-craclting zone 6. The reforming reactions are effected at a pressure of about 250 psig., a temperature of about 950 F., a liquid hourly space velocity of about 2.0 and a hydrogen/hydrocarbon mole ratio of about 70:10. The catalytic composite constitutes an alumina carrier material containing 0.55 percent by weight of platinum, 0.25 percent by weight of tin and 1.05 percent by weight of combined chloride, all of which are computed on an elemental basis. The reformed product effluent is withdrawn from reforming zone 12 by way of line Ml.

Not illustrated in the drawing is the technique whereby the reformed product effluent is separated to provide; (l) a hydrogen-rich gaseous phase for recycle to various zones of the process; (2) a propane/butane concentrate, for use as LPG or in alkylate production; and, (3) a pentane-plus fraction for direct use in motor fuel blending. in those instances where the aromatic concentrate is separately recovered from the pentaneplus portion, further separation is effected to recover a pentane/hexane concentrate, for conversion in the previously described hydroisomerization zone. The remaining heptane-plus portion continues through line 14 into solvent extraction zone 15 for aromatic separation from a paraffinic raffmate. The aromatics are withdrawn through line 116, and the paraffins, where such a technique is desired, are recycled, via lines 117 and 113, with the 390 F.-plus bottoms stream from the fractionator 8, for further reaction in l-cracking zone 6.

Considering only the pentane-plus portion of the product effluent, without the additional benefits of isomerization and alkylation, a motor fuel product is recovered having a clear research octane rating in excess of 100.0. This is accompanied by virtually negligible yield loss to methane, ethane and propane i.e., approximately 3.5 percent by weight. Since the fresh feed charge stock contained kerosene boiling range by drocarbons which were converted into lower molecular weight material, the volumetric yield exceeds 100.0 percent. Additionally, the clear gasoline pool comprises a greater volumetric percent of aromatic hydrocarbons.

We claim as our invention:

1. A process for the simultaneous production of an aromatic concentrate and isobutane, from a charge stock containing hydrocarbons boiling above a temperature of about 400 F., which process comprises the steps of:

a. separating said charge stock, in a first separation zone, to provide a first fraction having an end boiling point in the range of about 225 F. to about 260 F., a second fraction having an initial boiling point in the range of 225 F. to 260 F. and an end boiling point in the range of about 300 F. to about 410 F., and a third fraction having an initial boiling point in the range of about300 F. to about 410 F. and an end boiling point greater than about 410 F.;

b. reacting said first and third fractions with hydrogen, in a hydrocracking reaction zone, at hydrocracking conditions and in contact with a hydrocracking catalytic composite of a Group VIII noble metal component, and the reaction product of alumina and a gem-polyhalo compound;

c. separating the resulting hydrocracked product effluent, in a second separation zone, to provide a heptane-plus concentrate and to recover said isobutane;

d. reacting said heptane-plus concentrate and said second fraction in a catalytic reforming zone, at re forming conditions selected to convert naphthenic and paraffinic hydrocarbons to aromatic hydrocarbons; and,

e. recovering said aromatic concentrate from the resulting reformed product effluent.

2. The process of claim 11 further characterized in that said reformed product effluent is separated, in a third separation zone to provide a saturated paraffinic stream and to recover said aromatic concentrate.

3. The process of claim 2 further characterized in that said paraffinic stream is reacted with hydrogen in said hydrocracking reaction zone.

4. The process of claim 11 further characterized in that said hydrocracked product effluent is separated to provide a pentane/hexane concentrate and a propane/- butane concentrate.

5. The process of claim 4 further characterized in that said pentane/hexane concentrate is reacted with hydrogen in a hydroisomerization reaction zone, at isomerization conditions selected to produce pentane and hexane isomers.

6. The process of claim ll further characterized in that said hydrocracking catalytic composite comprises from about 0.01 percent to about 2.0 percent by weight of a platinum or palladium component.

7. The process of claim 11 further characterized in that said gem-polyhalo compound is a Friedel-Crafts metal halide.

bon, in an alkylation reaction zone, at alkylation conditions selected to produce a normally liquid alkylated hydrocarbon stream.

10. The process of claim 2 further characterized in that said third separation zone is a solvent extraction zone. 

2. The process of claim 1 further characterized in that said reformed product effluent is separated, in a third separation zone to provide a saturated paraffinic stream and to recover said aromatic concentrate.
 3. The process of claim 2 further characterized in that said paraffinic stream is reacted with hydrogen in said hydrocracking reaction zone.
 4. The process of claim 1 further characterized in that said hydrocracked product effluent is separated to provide a pentane/hexane concentrate and a propane/butane concentrate.
 5. The process of claim 4 further characterized in that said pentane/hexane concentrate is reacted with hydrogen in a hydroisomerization reaction zone, at isomerization conditions selected to produce pentane and hexane isomers.
 6. The process of claim 1 further characterized in that said hydrocracking catalytic composite comprises from about 0.01 percent to about 2.0 percent by weight of a platinum or palladium component.
 7. The process of claim 1 further characterized in that said gem-polyhalo compound is a Friedel-Crafts metal halide.
 8. The process of claim 1 further characterized in that said hydrocracking conditions include a maximum catalyst bed temperature of from 300* F. to about 600* F. and a pressure in the range of about 200 to about 500 psig.
 9. The process of claim 1 further characterized in that said isobutane is reacted with an olefinic hydrocarbon, in an alkylation reaction zone, at alkylation conditions selected to produce a normally liquid alkylated hydrocarbon stream.
 10. The process of claim 2 further characterized in that said third separation zone is a solvent extraction zone. 